KR20020087062A - Process for the preparation of propane-1,3-diol by vapor phase hydrogenation of 3-hydroxypropanal, beta-propiolactone, oligomers of beta-propiolactone, esters of 3-hydroxypropanoic acid or mixtures thereof - Google Patents

Abstract

Processes for the preparation of propane-1,3-diol are described. The process of the invention is substantially selected from hydrogen containing gases and 3-hydroxypropanols, β-propiolactones, oligomers of β-propiolactones, esters of 3-hydroxypropanoic acid and mixtures of two or more thereof Reacting the gaseous raw material mixture consisting of anhydrous feedstock under hydrogenation conditions in a hydrogenation zone in which a heterogeneous hydrogenation catalyst is present and recovering the reaction product comprising propane-1,3-diol It includes.

Description

Process for producing propane-1,3-diol by gas phase hydrogenation of 3-hydroxypropanol, beta-propiolactone, beta-propiolactone oligomer, ester of 3-hydroxypropanoic acid or mixtures thereof { PROCESS FOR THE PREPARATION OF PROPANE-1,3-DIOL BY VAPOR PHASE HYDROGENATION OF 3-HYDROXYPROPANAL, BETA-PROPIOLACTONE, OLIGOMERS OF BETA-PROPIOLACTONE, ESTERS OF 3-HYDROXYPROPANOIC ACID OR MIXTURES THEREOF}

Propane-1,3-diol is used as an intermediate in the preparation of polyesters for the production of fibers or films. Propane-1,3-diol may be prepared by a two-step process in which ethylene oxide is subjected to oxonation reaction according to the hydrogenation reaction of the following formula:

U.S. Pat.No. 5,981,808 uses non-phosphine-ligated cobalt compounds as oxonation catalysts in essentially non-water-miscible solvents, followed by extraction of water to produce the oxonation product. A process for separating the catalyst from the prepared 3-hydroxypropanal is described. This is followed by hydrogenation of the aqueous mixture containing 3-hydroxypropanal. U.S. Patent 5,585,528 proposes the addition of fatty affinity tertiary amines as accelerators of such processes. The use of methyl t-butyl ether to extract and reuse the cobalt catalyst by extracting the aqueous mixture is described in US Pat. No. 5,770,776. US Pat. No. 5,786,524 teaches a similar method using a rhodium catalyst as another catalyst in the osonation step.

However, a drawback of the process is that high levels of by-products are produced during liquid phase hydrogenation under recommended hydrogenation conditions, ie 220 ° C. and 100 bar (1000 kPa), which is an intermediate of 3-hydroxypropanal. Under such conditions, the conversion of 3-hydroxypropanal is only about 90% and up to 10% is converted to by-products.

It has also been proposed to combine the osonation step and the hydrogenation step in a one step process with minimal production of 3-hydroxypropanal as a by-product. Such a one-step process can be done using phosphine complexes of cobalt carbonyl as the main catalyst component. However, the use of ruthenium compounds as catalysts has also been proposed. An organic solvent is used to extract the water used to separate propane-1,3-diol from the osonation mixture in the reaction. Ethylene oxide conversion of 55% is reported with 87% selectivity to propane-1,3-diol.

US Pat. Nos. 5,310,948 and 5,359,081 teach the formation of β-propiolactone or polymers thereof by reacting carbon monoxide with ethylene oxide in the presence of a cobalt containing catalyst system comprising a source of cobalt and a hydroxy group substituted pyridine.

Alternatively, propane-1,3-diol can be produced from glycerol using recombinant bacteria expressing recombinant diol dehydratase. Such a method is taught in US Pat. No. 5,821,092.

Also proposed is a method of dehydrating acrolein to form 3-hydroxypropanal followed by hydrogenation. See US Patent No. 5,364,987 in this regard.

U.S. Patent No. 5,334,778 catalyzes the hydrogenation of 3-hydroxypropanal in aqueous solution at 50-95% 3-hydroxypropanal conversion in the presence of a hydrogenation catalyst at 30-80 캜, followed by hydrogenation. A process has been proposed to produce propane-1,3-diol with a residual carbonyl content of less than 500 ppm by continuing the reaction at 120 ° C. to 140 ° C. to achieve substantially 100% 3-hydroxypropanal conversion.

In general, however, both glycerol and icrolein are more expensive than ethylene oxide and are not readily available. Thus, the production of propane-1,3-diol by the last two methods mentioned above is not currently economical.

It is desirable to provide an improved method for the production of propane-1,3-diol. Hydrogenation of suitable carbonyl compounds also increases selectivity to propane-1,3-diol and reduces the amount of undesirable by-products such as propane-1-ol that cannot be converted immediately to propane-1,3-diol. It is desirable to provide a process for producing propane-1,3-diol. Also as intermediate compounds that can be made from ethylene oxide with minimal formation of undesirable by-products, it is possible to prepare 3-hydroxypropanols, β-propiolactones, oligomers of β-propiolactones or 3-hydroxypropionic acid. It is desirable to provide a method for producing propane-1,3-diol by hydrogenating an intermediate compound having at least one carbon-oxygen double bond such as an ester.

The present invention relates to the preparation of propane-1,3-diol.

It is therefore an object of the present invention to provide an improved process for the preparation of propane-1,3-diol. In another aspect, the present invention is to provide a method of using an optimized catalyst system in the hydrogenation step as a hydrogenation method for the production of propane-1,3-diol. Another object of the invention is to have at least one carbon-oxygen double bond, preferably 3-hydroxypropanal, β-propiolactone, oligomer of β-propiolactone, ester of 3-hydroxypropionic acid and A process for producing propane-1,3-diol by hydrogenation of an intermediate compound selected from two or more mixtures thereof, comprising: propan-1-ol, which cannot be readily converted to the desired propane-1,3-diol It is to provide a method of reducing the formation of such undesirable by-products.

According to the present invention, a method for preparing propane-1,3-diol, comprising a hydrogen-containing gas and an oligomer of 3-hydroxypropanal, β-propiolactone, β-propiolactone, and 3-hydroxypropanoic acid And forming a gaseous feed mixture consisting of a substantially anhydrous feedstock selected from a mixture of two or more thereof, containing a heterogeneous hydrogenation catalyst and hydrogenating the feedstock to propane- Feeding said gaseous feed mixture to a hydrogenation zone maintained under hydrogenation conditions capable of converting to 1,3-diol, and propane-1,3-diol from the hydrogenation zone; There is provided a method comprising recovering a reaction product comprising.

The raw material fed to the hydrogenation step is substantially anhydrous, that is to say that the moisture content is about 5% (w / v) or less, preferably about 1% (w / v) or less, and more preferably about 0.1% (w) / v) or less. The feedstock is selected from 3-hydroxypropanal, β-propiolactone, oligomers of β-propiolactone, esters of 3-hydroxypropanoic acid and mixtures of two or more thereof. β-propiolactone can cause spontaneous polymerization to form oligomers of β-propiolactone. The presence of small amounts of such oligomers in the feedstock fed to the hydrogenation zone is generally undesirable because of their relatively low volatility. Therefore, it is generally preferred to use a feedstock supplied to the hydrogenation zone in which the oligomer content of β-propiolactone is less than about 10 mol%. It is therefore usually desirable to minimize the proportion of oligomers of β-propiolactone in the feedstock fed to the hydrogenation zone.

The hydrogenation reaction is carried out using a gaseous raw material mixture which feeds to the hydrogenation zone, which contains a hydrogen containing gas in addition to the feedstock. The hydrogen containing gas is preferably substantially free of oxides of carbon, but may contain at least 50% v / v of one or more inert gases such as nitrogen, argon and helium, preferably about 10% v / or less, more preferably about 5% v / v, ie about 1% v / v or less.

The conditions of the hydrogenation reaction may be chosen such that the reaction mixture exiting the hydrogenation zone is also in the gas phase. Alternatively, however, it may not only be possible to utilize hydrogenation conditions in which the reaction mixture at the outlet end of the hydrogenation zone is below the dew point and as a result at least some of its condensable components are present in the liquid phase. have.

The hydrogenation zone is conveniently equipped with a fixed bed of granular hydrogenation catalyst. The hydrogenation zone can be provided with one or more catalyst beds if desired and the hydrogenation catalyst on one catalyst can be different from the hydrogenation catalyst on at least another catalyst if desired.

The hydrogenation catalyst is preferably selected from noble metal catalysts such as palladium catalysts and copper containing catalysts. More preferably, the hydrogenation catalyst is a copper containing catalyst. The active catalysts in the hydrogenation catalyst are two or more of these, such as chromia, zinc oxide, alumina, silica, silica-alumina, silicon carbide, zirconia, titania, carbon or a mixture of chromia and carbon, for example. At least partially supported on a support selected from a mixture of:

Examples of suitable copper-containing catalysts that may be utilized in the process of the present invention include reduced copper oxide / zinc oxide hydrogenation catalysts, reduced manganese-promoted copper catalysts, reduced copper chromite catalysts and reduction promoted ( reduced promoted) copper chromite catalysts.

The hydrogenation catalyst is preferably a reduced manganese-promoted copper catalyst. The total surface area of such manganese-promoted copper catalysts is preferably at least about 15 m 2 / g, more preferably at least 20 m 2 / g, even more preferably at least 25 m 2 / g in unreduced form.

A particularly preferred hydrogenation catalyst is a reduced manganese-promoted copper catalyst available such as the DRD92 / 89A catalyst from Kvaerner Process Technology Limited of The Technology Center, TS17 6PY Stockton-on-Tees-Sonby Princeton Drive, UK. An alternative is a reduced manganese-promoted copper catalyst available such as DRD92 / 89B from Kvaerner Process Technology Limited.

The hydrogenation step is carried out under gaseous raw material feed conditions such that the feed stream entering the hydrogenation zone is above the dew point and thus the gaseous feed stream. The reaction product mixture from the hydrogenation zone may be recovered at a temperature above the dew point such that it is also gaseous, or at a temperature below the dew point such that at least some of the condensable components of the mixture are liquid.

Although the hydrogenation process of the present invention may be carried out in a tubular reactor under substantially isothermal conditions, the adiabatic reactor can be manufactured much cheaper than a tubular reactor and is generally insulated using a fixed catalyst bed. It would be desirable to operate under hydrogenation conditions. However, in order to keep the temperature at which the reaction mixture is exposed to the desired limits in the plant design, such as the selection of particularly suitable gas: feedstock ratios and catalyst size, the temperature rise may be at or below a reasonable value, generally about 20, in any catalyst bed position. Care must be taken to limit the temperature to less than or equal to ℃. This may limit the formation of byproduct 1-propanol.

It will generally be desirable for the molar ratio of hydrogen containing gas to feedstock in the gaseous feed stream to the hydrogenation zone to be in the range of about 50: 1 to about 1000: 1.

In general, the temperature of the feed to the hydrogenation zone is from about 130 ° C. to about 180 ° C., more preferably from about 135 ° C. to about 150 ° C., and the pressure of the feed to the hydrogenation zone is about 50 psia (about 344.74 kPa) to about 2000 psia (about 13789.52 kPa), for example about 350 psia (about 2413.17 kPa) to about 1000 psia (about 6894.76 kPa). The feedstock is also preferably fed to the first hydrogenation zone at a rate corresponding to the hourly space velocity of the liquid from about 0.05 to about 5.0 h ⁻¹ . The unreacted hydrogen containing gas is preferably recycled for later use.

If desired, the feedstock to the hydrogenation zone can be diluted with a solvent, such as methanol, which is stable under the hydrogenation conditions utilized.

The invention is illustrated more specifically in the following examples.

Example 1

This example is intended to provide a simulation of conditions similar to those that may exist in a commercial hydrogenation reactor for hydrogenating methyl 3-hydroxypropionate utilizing recycled hydrogen saturated with a methanol co-product. It is.

A solution of the methyl 3-hydroxypropionate crude product (about 90% purity) was diluted with approximately equal weight methanol to make a stock solution. This stock solution was subjected to gas phase hydrogenation in a once-through adiabatic fixed-bed reactor. The reactor was made of a 95 cm long, 20.96 mm inner tube with an oil jacket to reduce heat loss. The reactor was charged with 100 ml of DRD92 / 89A catalyst, available from Kvaerner Process Technology Limited of The Technology Center, TS17 6PY Stockton-on-Tees-Sonby Princeton Drive, UK. The catalyst was reduced by the same procedure as described in US Pat. No. 5,030,609.

The stock solution was fed to a heater at a feed rate of 12 ml / h and evaporated by a pure hydrogen stream of 1000 Nl / h (ie 760 mmHg [101.33 kPa] and 1000 liters per hour as measured at 0 ° C). The gas mixture was passed over the catalyst at a pressure of 400 psig (gauge pressure 2757.90 kPa) and at a temperature of 138 ° C. The reaction product mixture present in the reactor was cooled down and the condensate was collected. Raw materials and products were analyzed by gas chromatography using a CP SIL 19 capillary column 60 meters with a film thickness of 1.3 μm.

The conversion of methyl 3-hydroxypropionate was 73.3%, the selectivity to propane-1,3-diol was determined to be 83.4%, the selectivity to 1-propanol was 5.8%. By-products can be recycled because they are believed to contain substances that can be converted into propane-1,3-diol upon hydrogenation.

Example 2

The general procedure of Example 1 was repeated except that the temperature was changed to 148 ° C. The conversion of methyl 3-hydroxypropionate was determined to be 84.9%, 84.8% selectivity to propane-1,3-diol, and 7.2% selectivity to 1-propanol. This slightly higher temperature operation provides the expected increase in conversion with a slight increase in selectivity to 1-propanol.

Example 3

The general procedure of Example 1 was repeated except that the temperature was changed to 174 ° C. The conversion rate of methyl 3-hydroxypropionate was determined to be 99.95%, 13.1% selectivity to propane-1,3-diol, and 81.8% selectivity to 1-propanol. This indicates that manipulation at high temperatures promotes the formation of 1-propanol, which is an alcohol, rather than the formation of propane-1,3-diol.

Example 4

The same crude solution as used in Example 1 was subjected to gas phase hydrogenation in an adiabatic fixed bed reactor system where the excess gas was recycled after condensation of the reactor product stream. Replenishing pure hydrogen was fed to the system to maintain a constant pressure in the system. The reactor was made of a tube of 200 cm in length and 26.64 mm in inner diameter with insulated, electrically traced heating means to prevent net heat loss. The reactor received a 250 ml dose of DRD92 / 89A catalyst. The catalyst was reduced by the same procedure as described in US Pat. No. 5,030,609.

The raw material solution was fed to the heater at a feed rate of 80 ml / h and evaporated by a mixed gas stream of recycle gas and pure hydrogen at a flow rate of 10900 Nl / h. The gas mixture was passed over the catalyst at a pressure of 885 psig (gauge pressure 6101.86 kPa) and an inlet temperature of 149 ° C. The outlet temperature was measured at 150 ° C. The reactor resulting mixture was cooled and the condensate was collected. Raw materials and products were analyzed by gas chromatography using a CP SIL 19 capillary column 60 meters with a film thickness of 1.3 μm.

The conversion of methyl 3-hydroxypropionate was 77.8%, the selectivity to propane-1,3-diol was determined to be 78.9%, the selectivity to 1-propanol was 14.8%.

Example 5

The general procedure of Example 4 was repeated except that the inlet pressure was 735 psig (5067.65 kPa) with a mixed gas flow rate of recycle gas and supplemental pure hydrogen of 8116 Nl / h. The conversion of methyl 3-hydroxypropionate was 61.1%, the selectivity to propane-1,3-diol was determined to be 79.0%, the selectivity to 1-propanol was 14.4%. Operation at these low pressures and low gas velocities leads to a reduction in conversion as expected but the selectivity to propane-1,3-diol and 1-propanol is similar.

Example 6

A solution of methyl 3-hydroxypropionate crude (pure about 98%) was diluted with approximately equal weight methanol to make a stock solution. This raw material solution was subjected to gas phase hydrogenation in the apparatus used in Example 4.

The raw material solution was fed to the heater at a feed rate of 80.8 ml / h and evaporated by a mixed gas stream of recycle gas and pure hydrogen at a flow rate of 8717 Nl / h. The gas mixture was passed over the catalyst at a pressure of 885 psig (gauge pressure 6101.86 kPa) and an inlet temperature of 148 ° C. The outlet temperature was measured at 149 ° C.

The reactor resulting mixture was cooled and the condensate was collected. Raw materials and products were analyzed by the method used in Example 4.

The conversion of methyl 3-hydroxypropionate was 71.3%, the selectivity to propane-1,3-diol was determined to be 80.6% and the selectivity to 1-propanol was 12.5%. By-products can be recycled because they are believed to contain substances that can be converted into propane-1,3-diol upon hydrogenation.

Example 7

The general procedure of Example 6 was repeated except that the feed flow rate was 61.4 ml / h, the recycle airflow was 6537 Nl / h, and the inlet temperature was 149 ° C. The outlet temperature was measured at 150 ° C.

The conversion of methyl 3-hydroxypropionate was 84.1%, the selectivity to propane-1,3-diol was 73.7%, and the selectivity to 1-propanol was 21.5%.

Example 8

The general procedure of Example 7 was repeated except that the inlet temperature was 154 ° C. The outlet temperature was measured at 155 ° C.

The conversion of methyl 3-hydroxypropionate was 87.3%, the selectivity to propane-1,3-diol was determined to be 75.2%, the selectivity to 1-propanol was 18.7%.

Example 9

The general procedure of Example 8 was repeated except that the flow rate of the mixed supplemental air stream of the recycle gas and pure hydrogen was 5226 Nl / h and the inlet temperature was 152 ° C. The outlet temperature was measured at 153 ° C.

The conversion of methyl 3-hydroxypropionate was 86.7%, the selectivity to propane-1,3-diol was determined to be 76.2%, the selectivity to 1-propanol was 16.9%.

Claims (13)

As a method for preparing propane-1,3-diol, Hydrogen containing gas and substantially anhydrous feedstock selected from 3-hydroxypropanal, β-propiolactone, oligomers of β-propiolactone, esters of 3-hydroxypropanoic acid and mixtures of two or more thereof Forming a vaporous feed mixture comprising the same, holding under heterogeneous hydrogenation catalyst and hydrogenating the feedstock to convert it to propane-1,3-diol Propane-1 comprising supplying the gaseous crude mixture to a hydrogenation zone, and recovering the reaction product comprising propane-1,3-diol from the hydrogenation zone; Process for the preparation of 3-diols. The method of claim 1, A process for producing propane-1,3-diol wherein the feedstock comprises an alkyl ester or a hydroxyalkyl ester of 3-hydroxypropanoic acid. The method according to claim 1 or 2, A process for producing propane-1,3-diol, wherein the hydrogenation reaction is carried out using a fixed bed of granular hydrogenation catalyst. The method according to any one of claims 1 to 3, And said hydrogenation catalyst is selected from a noble metal catalyst and a copper containing catalyst. The method of claim 4, wherein A method for producing propane-1,3-diol, wherein the hydrogenation catalyst is a copper-containing catalyst. The method of claim 5 Propane-1, wherein the copper containing catalyst is selected from a reduced copper oxide / zinc oxide hydrogenation catalyst, a reduced manganese-promoted copper catalyst, a reduced copper chromite catalyst and a reduced promoted copper chromite catalyst, Process for the preparation of 3-diols. The method according to claim 5 or 6, A method for producing propane-1,3-diol, wherein the copper-containing catalyst is a reduced manganese-promoted copper catalyst. The method according to any one of claims 1 to 7, A process for producing propane-1,3-diol wherein the molar ratio of hydrogen-containing gas to feedstock in a gaseous crude mixture is in the range of about 50: 1 to about 1000: 1. The method according to any one of claims 1 to 8, A process for preparing propane-1,3-diol, wherein the feed temperature to the hydrogenation zone is from about 130 ° C to about 180 ° C. The method of claim 9, A process for preparing propane-1,3-diol, wherein the feed temperature to the hydrogenation zone is from about 135 ° C to about 150 ° C. The method according to any one of claims 1 to 10, A process for preparing propane-1,3-diol, wherein the feed pressure to the hydrogenation zone is from about 50 psia (about 344.74 kPa) to about 2000 psia (about 13789.52 kPa). The method of claim 11, A process for preparing propane-1,3-diol, wherein the feed pressure to the hydrogenation zone is from about 350 psia (about 2413.17 kPa) to about 1000 psia (about 6894.76 kPa). The method according to any one of claims 1 to 12, Wherein said feedstock is fed to said first hydrogenation zone at a rate corresponding to an hourly space velocity of a liquid of from about 0.05 to about 5.0 h ⁻¹ .